Note: Descriptions are shown in the official language in which they were submitted.
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
LOW-PRESSURE FISCHER-TROPSCH PROCESS
Fidd of the Invention
This invention relates generally to a. low pressure Fischer-Tropsch process
for
converting carbon monoxide and hydrogen to diesel fuel or diesel blending
stock.
naeltEround of the Invention
The Fischer-Tropsch (FT) process for converting carbon monoxide and
hydrogen to liquid motor fuels and/or wax has been known since the 1920's.
During the Second World War synthetic diesel was manufactured in Germany
using coal gasification to supply a 1:1 ratio of hydrogen and carbon monoxide
for
conversion to fuel hydrocarbons. Because of trade sanctions and the paucity of
natural gas, South Africa further developed the coal via gasification route to
synthesis
gas and employed a fixed-bed iron Fischer-Tropsch catalyst. Iron catalysts are
very
active for the water-gas shift reaction which moves the gas composition from a
deficiency of hydrogen and closer to the optimum 112/C0 ratio of around 2Ø
When
large natural gas supplies were developed, steam and autothermal reformers
were
employed to produce the synthesis gas feedstock to slurry-bed FT reactors
using
cobalt or iron catalysts.
In Gas-To-Liquids (GTL) plants, compromises must be made between liquid
product yield and plant operating and capital costs. For example, if there is
a market
for electricity, a step= reformer design may be chosen because this technology
produces a large amount of waste heat: flue gas heat can be converted to
electricity
using an 'economiser' and steam turbine. If conservation of natural gas
feedstock and
low capital cost are paramount, autothermal or partial oxidation reformers
using air
are favored.
1
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
Another factor in selecting the best reformer type is the nature of the
reformer
hydrocarbon feed gas. If the gas is rich in CO2, this can be advantageous
because the
desired 1-12/C0 ratio can then be achieved directly in the reformer gas
without the need
to remove excess hydrogen, and some of the CO2 is converted to CO, increasing
the
potential volume of liquid hydrocarbon product that can be produced.
Additionally,
the volume of steam that is required is reduced, which reduces the process
energy
requirements,
The market for Fischer-Tropsch (FT) processes is concentrated on large
"World-Scale" plants with natural gas feed rates of greater than 200 million
scfd
because of the considerable economies of scale. These plants operate at high-
pressure,
about 450 psia, and use extensive recycling of tail gas in the FT reactor.
For, example,
the Norsk Hydro plant design has a recycle ratio of about 3Ø The emphasis is
on
achieving the maximum wax yield, In terms of product slate, these large plants
strive
for the maximum yield of FT waxes in order to minimize the formation of C1-05
products. The waxes are then hycirocracked to primarily diesel and naphtha
fractions.
Unfortunately, light hydrocarbons are also formed in this process. The
reformers
usually use some form of autothermal reforming with oxygen that is produced
cryogenically from air, an expensive process in terms of operating cost and
capital
cost. The economies of scale justify the use of high operating pressure, the
use of
oxygen natural gas reforming, extensive tail gas recycling to the FT reactor
for
increasing synthesis gas conversion and controlling heat removal and product
wax
hydrocracIdng. To date, an economical FT plant design has not been developed
for
small plants with capacities of less than 100 million scfd,
The catalytic hydrogenation of carbon monoxide to produce a variety of
products ranging from methane to heavy hydrocarbons (up to Cg o and higher) as
well
as oxygenated hydrocarbons is usually referred to as Fischer-Tropsch
synthesis. The
high molecular weight hydrocarbon product primarily comprises normal paraffins
which cannot be used directly as motor fuels because their cold properties are
not
compatible. After further hydroprocessing, Fischer-Tropsch hydrocarbon
products can
be transformed into products with a higher added value such as diesel, jet
fuel or
2
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
kerosene. Consequently, it is desirable to maximize the production of high
value
liquid hydrocarbons directly so that component separation or hydrocracking are
not
necessary.
Catalytically active group VIII, in particular, iron, cobalt and nickel are
used
as Fischer-Tropsch catalysts; cobalt/ruthenium is one of the most common
catalyzing
systems. Further, the catalyst usually contains a support or carrier metal as
well as a
promoter, e.g., rhenium.
Summary of the Invention
According to one aspect of the invention, there is provided Fischer-Trope.ch
(FT) process having a cobalt catalyst with crystallites, wherein the
crystallites have an
average diameter greater that 16 nanometers. The process produces a liquid
hydrocarbon product containing leas than 10 weight percent wax (>C23) and
greater
than 65% diesel (C9-C23). The process can have a FT catalyst support for the
cobalt
catalyst, wherein the catalyst support is selected from the group of catalyst
supports
consisting of alumina, zirconia, titania and silica. The cobalt catalyst can
have a
catalyst loading that is greater than 10 weight A. The operating pressure for
the
Fischer-Tropsch process can be less than 200 pain. Promoters can be utilized
in this
process, in which case the promotors are selected from the group consisting
of:
ruthenium, rhenium, rhodium, nickel, zirconium, titanium, and mixtures
thereof. A
flash distillation can be conducted on the process to reduce the naphtha cut.
The
process can use a FT reactor that does not use tailgas recycle. The process
can also
use a reformer that uses air as an oxygen source. The reactor can be a fixed-
bed FT
reactor or a slurry bubble bed FT reactor.
According to another aspect of the invention, there is provided a FT process
operating at less than 200 psia, using an air autothermal reformer, and having
a CC)
conversion of at least 65 % and providing a diesel yield greater than 60% by
weight in
a single- pass FT reactor using a cobalt catalyst. The catalyst has a metallic
cobalt
3
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
loading greater than 5% by weight and rhenium loading of less than 2% by
weight on
a catalyst support material selected from the group of catalyst support
materials
comprising alumina, zirconia, and silica. The cobalt catalyst is in the form
of
crystallites, wherein the crystallites have an average diameter greater that
16
nanometers. The FT catalyst support material can be comprised of alumina. This
process can have a feed gas, wherein selective membranes or molecular sieves
are
employed to remove hydrogen from the feed gas. Alternatively, the cobalt
catalyst
loading can be greater than 6 weight % and operating pressure can be less than
100
psia. The reactor can further have a promoter, wherein said promoter comprises
a
promoter selected from the group of promoters consisting of ruthenium and
rhenium
or mixtures thereof.
According to yet another aspect of the invention, there is provided a FT
process operating at less than 200 psia, using an oxygen autotherrnal
reformer, and
having a CO conversion of at least 65 % and providing a diesel yield greater
than 60%
by weight in a FT reactor using a cobalt catalyst. The catalyst has a metallic
cobalt
loading greater than 5% by weight and a rhenium loading of less than 2% by
weight
on a catalyst support selected from the group of catalyst supports comprised
of
alumina, zirconia and silica materials. The cobalt catalyst is in the form of
crystallites,
said crystallites having an average diameter greater that 16 nanometers. The
FT
catalyst support can be comprised of alumina. The process can include a
tailgas from
the reformer, wherein the taigas is partially recycled to the reformer. The
process can
also include a feed gas wherein selective membranes or molecular sieves are
employed to remove hydrogen from the gas. Alternatively, the cobalt catalyst
loading
can be greater than 6 weight % and the operating pressure can be less than 100
psia.
The reactor can further have a promoter, wherein said promoter comprises a
promoter
selected from the group of 'promoters consisting of ruthenium or rhenium, or
mixtures thereof
According to yet another aspect of the invention, there is provided a FT
process operating at less than 200 psia, using an oxygen steam reformer, and
having a
CO conversion of at least 65 % and providing a diesel yield greater than 60%
in by
4
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
weight in a FT reactor using a cobalt catalyst with a metallic cobalt loading
greater
than 5% by weight and rhenium loading of less than 2% by weight on a catalyst
support selected from the group of catalyst supports comprised of alumina,
zirconia,
or silica materials, or mixtures thereof. The cobalt catalyst is in the form
of
crystallites, wherein the crystallites have an average diameter greater that
16
nanometers. The FT catalyst support can be comprised of alumina. The process
can
include a feed gas, wherein selective membranes or molecular sieves are
employed to
remove hydrogen from the feed gas. The process can further include a tailgas
from
the reformer, wherein some or all of the tailgas is burned to provide heat to
the
reformer. Alternatively, the cobalt catalyst loading can be greater than 6
weight %
and the operating pressure can be less than 100 psis. The reactor can further
have a
promoter, wherein said promoter comprises a promoter selected from the group
of
promoters consisting of ruthenium or rhenium, or mixtures thereof.
According to yet another aspect of the invention, there is provided a FT
process operating at less than 200 psia, using an air or oxygen partial
oxidation
reformer, and having a CO conversion of greater than 65 % and providing a
diesel
yield greater than 60% by weight in a FT reactor using a cobalt catalyst with
a
metallic cobalt loading greater than 5% by weight and rhenium loading of less
than
2% by weight on a FT catalyst support selected from the group of catalyst
supports
comprising alumina, zirconia, and silica materials. The cobalt catalyst is in
the form
of crystallites, and the crystallites have an average diameter greater that 16
nanometers, The FT catalyst support can be comprised of alumina. The process
can
include a feed gas, wherein selective membranes or molecular sieves are
employed to
remove hydrogen from the feed gas. The process can further include a tailgas
from
the reformer, wherein some or all of the tailgas is burned to provide heat to
the
reformer. Alternatively, the cobalt catalyst loading can be greater than 6
weight %
and the operating pressure can be less than 100 psia. The reactor can further
have a
promoter, wherein said promoter comprises a promoter selected from the group
of
promoters consisting of ruthenium or rhenium, or mixtures thereof.
5
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
Brief Deserintion of the Drawings
Figure 1 is a process flow diagram for a particular embodiment of the
invention;
Figure 2 is a flow diagram for flash separation of naphtha and diesel
hydrocarbon
fractions as a subsequent step to the Fischer-Tropsch process;
Figure 3 is a graph showing C5+ carbon number distribution for the catalyst of
Example 3 (trilobes) at 190 C;
Figure 4 is a graph showing the effect of pressure on the performance of the
catalyst
of Example 4;
Figure 5 is a graph of C5+ carbon number distribution of the catalyst of
Example 7 at
190 C, 70 psia;
Figure 6 is a graph of the C5+ carbon number distribution for the catalyst of
Example
8a (LD-5) at 200 C, 70 psia;
Figure 7 is a graph of the C5+ carbon number distribution for the catalyst of
Example
9 (F-220) at 190 C., 70 psia
Figure 8 is a graph of the C5+ carbon number distribution for the catalyst of
Example
10 using a Ruthenium promoter;
Figure 9 is a graph of the C5+ carbon number distribution for Catalyst of
Example
11 (Aerolyst 3038 silica);
Figure 10 is a graph showing the relationship of cobalt catalyst crystallite
size to wax
content of a C5+ FT product; and
Figure 11 is a graph showing a comparison of catalyst used in Example 9 carbon
distribution with a traditional Anderson-Shultz-Flory distribution.
6
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
In all Figures showing graphs of carbon numbers, naphtha is indicated by
large squares, diesel by diamonds and light waxes by small squares.
petailed Description of Embodiments of the Invention
JAtroduction
In Fischer-Tropsch processes, various parameters such as the size and shape of
cobalt crystallites affect the activity of cobalt supported catalysts. The
size of metal
crystallites controls the number of active sites available for reduction
(dispersion) and
degree of reduction.
Under certain pretreatment and activation conditions, a strong interaction
between cobalt metal and oxide supports forms undesirable cobalt-support
structures,
for example, cobalt aluminate, which may require high reduction temperature.
High
reduction temperature can result in sintering cobalt crystallites and forming
large
cobalt metal clusters. Not only temperature treatments, but also cobalt metal
precursors and metal loading, as well as metal promoters affect the size of
cobalt
crystallites. Low cobalt metal loading could result in high metal dispersion
and small
crystallites but enhances the metal-support interaction leading to poor
reducibility and
low catalyst activity.
Hydrogenation of carbon monoxide using cobalt-supported catalyst is directly
proportional to the amount of exposed cobalt atoms. Therefore, increasing
cobalt
metal dispersion on the oxide support surface, logically, enhances the
catalyst activity
and C5+ selectivity. However, small cobalt crystallites strongly interact with
the
oxide support forming unreclucible cobalt-support systems. The strong
correlation
between cobalt metal crystallites and reducibility influences the catalyst
activity and
may produce undesirable products. Under typical Fischer-Tropsch reaction
conditions cobalt crystallite size range (9-200 rim) and dispersion range (11-
0.5 %)
have minor influence on C5+ selectivity. Nevertheless, smaller cobalt
crystallites
7
CA 02748216 2015-08-18
suffer from serious deactivation. In fact, Barbie et. al 2000 studied the
correlation
between the deactivation rate and cobalt crystallite size and observed a peak
at 5.5 nm.
Embodiments Of the Invention
Embodiments of the invention described herein relate to a low-pressure
Fischer-Tropsch process and a catalyst that produces a high diesel-fraction
yield.
Process pressure can be below 200 psia. The catalyst is cobalt deposited at
greater
than 5 weight percent on gamma alumina, optionally along with rhenium or
ruthenium at 0.01 -2 wt. %, and have crystallites having an average diameter
greater
than 16 nanometers. It has been discovered that this catalyst is very
effective at low
pressures in converting synthesis gas into diesel in high yield, producing a
liquid
hydrocarbon product containing less than 10 wt.% wax (>C23) and greater than
65%
diesel (C9-C23). The present embodiments are particularly well suited to
conversion
of low pressure gases containing low molecular weight hydrocarbons into FT
liquids.
Examples of applications are landfill gas, oil field solution gas and low
pressure gas
from de-pressured gas fields. In all these cases, multiple-stage gas and air
compression would be required in traditional FT plants. The high efficiency of
the
present FT catalyst enables high CO conversion and produces a product stream
containing up to 90+ wt. % diesel in a single pass. The use of air in the
natural gas
reformer provides a synthesis gas containing approximately 50% nitrogen, which
facilitates heat removal in the FT reactor as sensible heat and increases gas
velocity
and heat transfer efficiency, so that tail gas recycling is not needed.
Naphtha can be
partially separated from the hydrocarbon product by flash distillation at low
cost to
generate a more pure diesel product. This also serves to provide some product
cooling.
The liquid hydrocarbon product is excellent for blending with petroleum diesel
to
increase cetane number and reduce sulfur content.
The present embodiments can be applied to world-scale gas-to-liquid plants,
but also to small FT plants using less than 100 million scfd. When applied to
small
FT plants, the present embodiments strive for optimized economics with an
emphasis
on simplicity and minimized capital cost, possibly at the expense of
efficiency. The
- 8 -
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
following is a comparison of existing FT technologies compared to the present
embodiment as applied to small FT plants:
Existing FT Technologies Present Embodiments
Large plants, > 25MMscfd Small plants, < I OOMMscfd
High pressure, > 200 psia Low pressure, <200 psia
Oxygen to reformer Air to reformer
Extensive recycling to FT reactor or reformer No recycling ("once-through"
process)
Low single-pass FT CO conversion (<50%) High. single-pass conversion (>
65%)
Deliberate and extensive wax formation Less than 10% wax formation
Hydrocracking waxes No hydrocracicing operations
Multiple-pass FT reactors Single- pass-FT reactor
Low FT diesel yield (<50 %) High diesel yield (55-90 % of
hydrocarbon liquid)
In order to operate the FT process at high conversions with oxygen- blown
reformer synthesis gas, the approach has been to recycle tail gas in a high
proportion-
al at a ratio of 3.0 or greater based on fresh gas feed. A secondary
benefit is that the
fresh gas is diluted in carbon monoxide, which reduces the required rate of
heat
removal from the FT reactor, reduces hot-spotting and improves the product
slate.
However, tailgas recycling is a very energy and capital intensive activity.
The
separation of oxygen from air is also an energy and capital intensive
activity.
The approach taken in the present process is to use air in the reformer, which
gives a synthesis gas containing approximately 50 % nitrogen as inert diluent,
eliminating the need for tail gas recycling to moderate FT reactor heat
removal
requirements. Others employing air- blown synthesis gas in FT processes have
achieved the desired high CO conversions by using multiple FT reactors in
series,
which entails high capital costs and complex operation. The present process
achieves
9
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
high CO conversion in a simple single pass and a high diesel out by using a
special
catalyst as more particularly described below.
The catalyst in one embodiment employs an alumina support with high cobalt
concentration, along with a low level of rhenium to facilitate catalyst
reduction. The
high cobalt concentrations increase catalyst activity, enabling high single-
pass
synthesis gas conversion. The catalyst is made to have a relatively large
average
cobalt crystallite size and this gives selectivity to a substantially diesel
product.
The Anderson-Shultz-Florey theory predicts the FT hydrocarbons to cover a
very wide range of carbon numbers, from 1-60, whereas the most desirable
product is
diesel fuel (Co-C23, Chevron definition). In order to reduce the 'losses' of
CO to
making C1-05 hydrocarbons, a common approach is to strive to make mostly wax
in
the FT reactor and then, in a separate operation, to hydrocrack the wax to
mostly
diesel and naphtha.. Surprisingly, the process and catalyst of the present
embodiments
make diesel in high yield (to 90 wt%) directly in the FT reactor, obviating
the need
for expensive and complex hydrocracking
Because of the elimination of oxygen purification, high-pressure compression,
tail gas recycling and hydrocracking, the present process can be applied
economically
in much smaller plants than hitherto considered possible for FT technology.
Figure 1 shows the process flow diagram for the FT process of the present
embodiment, wherein the letters A-IC signify the following:
A Raw hydrocarbon-containing gas
= Hydrocarbon gas conditioning equipment
C Reformer
D Water
E Oxidizing gas
= Cooler
G Separator
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
H hydrogen removal (optional)
Fischer Tropsch reactor
Back-pressure controller
K Product cooling and recovery (2-options)
Letter A represents the raw hydrocarbon-containing process feed gas. This
could be from a wide variety of sources: for example, from a natural gas
field, a land-
fill facility (biogenic gas), a petroleum oil processing facility (solution
gas), among
others, The pressure of the gas for the present process can vary widely, from
to atmospheric pressure to 200 psia or higher. Single-stage or two-stage
compression
may be required, depending on the source pressure and the desired process
operating
pressure. For example, for landfill gas, the pressure is typically close to
atmospheric
pressure and blowers are used to transmit the gas into combustion equipment.
Solution gas, which is normally flared, must also be compressed to the process
operating pressure. There are also many old exploited and late-life natural
gas fields
with pressure too low for acceptance into pipelines that could make possible
feedstock
for the present process. Other natural gas sources, which may or may not be
stranded
(no access to a pipeline) may already be at or above the desired process
operation
pressure and these are also candidates. Another candidate is natural gas that
is too
high in inerts such as nitrogen to meet pipeline specifications.
Letter B represents hydrocarbon gas conditioning equipment. The gas may
require clean-up to remove components that would damage reformer or FT
catalyst.
Examples of these are mercury, hydrogen sulfide, silicones and organic
chlorides.
Organic chlorides, such as found in land-fill gas, produce hydrochloric acid
in the
reformer, which can cause severe corrosion. Silicones form a continuous
silicon
dioxide coating on the catalyst, blocking pores. Hydrogen sulphide is a
powerful FT
catalyst poison and is usually removed to 1.0 ppm or lower. Some gas, from
sweet-
gas fields, may not require any conditioning (clean-up).
The hydrocarbon concentration in the raw gas affects the economics of the
process because less hydrocarbon product is formed from the same volume of
feed
11
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
gas. Nevertheless, the process can operate with 50% or lower methane
concentration,
for example, using land-fill gas. There may even be reasons to operate the
process
even at a financial loss: for example to meet greenhouse gas government or
corporate
emission standards. The process can operate with feed gases containing only
methane
hydrocarbon or containing natural gas liquids by the application of known
reformer
technologies. The presence of carbon dioxide in the feed gas is beneficial.
Letter C represents the reformer, which may be of several types depending on
the composition of the feed gas. A signiAcant benefit of low pressure reformer
to operation is the lower rate of the Brouard reaction and diminution of
metal dusting.
Partial oxidation reformers normally operate at very high pressure i.e. 450
psia
or greater, and so are not optimum for a low-pressure FT process. It is
energetically
inefficient, and can easily make soot, however, it does not require water, and
makes a
ts syngas with a H2/C0 ratio near 2.0, optimum for FT catalysts. Partial
oxidation
reformers may be employed in the present process.
Steam reformers are capital expensive and require flue gas heat recovery to
maximize efficiency in large plants. Because the synthesis gas contains
relatively low
20 levels of inerts such as nitrogen, temperature control in the Fr reactor
can be difficult
without tail gas recycling to the FT reactor. However, the low level of inerts
enables
recycling of some tail gas to the reformer tube-side, supplementing natural
gas feed,
or to the shell side to provide heat. Keeping in mind that FT tail gas must be
combusted before venting in any event, this energy can be used for electrical
25 generation or, better yet, to provide the reformer heat which would be
otherwise be
provided from burning natural gas. For small FT plants, steam reformers are a
viable
choice. Steam reformers may be employed in the present process.
Autothermal reforming is an efficient process of relatively low capital cost
30 that uses moderate temperatures and modest steam concentrations to
produce a soot-
free synthesis gas with H2/C0 around 2.5 using low-0O2 natural gas feed, which
is
closer to the desired ratio than for steam reforming. However some hydrogen
removal
12
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
is still required for most natural gas feeds. If the feed gas contains greater
than about
33 % CO2, as is the case with land-fill gas feed, then an 112/C0 ratio of 2.0
can be
achieved without any recycle streams, and the water use can also be
diminished. This
is the most desired type of reformer for the present low-pressure FT
processes.
Letter D represents the optional water that is injected as steam into the
reformer. All reformer technologies except partial oxidation require the
injection of
steam.
Letter E represents an oxidizing gas, which could be air, oxygen or oxygen-
enriched air,
Letter F represents a cooler for reducing the reformer outlet temperature from
greater than 700 C. to close to ambient. The cooling may be done in several
stages,
but preferably in a single stage. The cooling may be achieved with shell- and-
tube or
plate- and. frame heat exchangers and the recovered energy may be utilized to
pre-
heat the reformer feed gases, as is well known in the industry. Another way of
cooling the reformer tail gas is by direct injection of water into the stream
or by
passing the stream through water in a vessel.
Letter G represents a separator for separating the reformer synthesis gas from
condensed water, so as to minimize the amount of water entering downstream
equipment,
Letter H represents optional hydrogen removal equipment such as Prism
hydrogen-selective membranes which are sold by Air Products, or Cynara
membranes
from Natco.
Certain reformer processes produce a synthesis gas too rich in hydrogen, some
of which must be removed to achieve optimum FT reactor performance. An ideal
H2/CO ratio is 2.0-2.1, whereas the raw synthesis gas may have a ratio of 3.0
or
13
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
higher. High hydrogen concentrations give rise to larger CO loss to producing
methane instead of the desired motor fuels or motor fuel precursor such as
naphtha.
Letter I represents typical FT reactors, which are of the fixed- bed or slurry
bubble type and either may be used. However, the fixed-bed is preferred in
small
plants because of its simplicity of operation and ease of scale-up.
Letter J represents a back-pressure controller which sets the process
pressure.
It may be placed in other locations depending on the product recovery and
possible
partial separation process employed.
Letter K represents product cooling and recovery. Product cooling is typically
accomplished by heat exchange with cold water and serves to pre-heat the water
for
use elsewhere in the FT plant. Separation is accomplished in a separator
vessel
designed for oil/water separation. However a second alternative is to flash-
cool the
FT reactor product before the aforementioned cooler-separator as shown in
Figure 2.
This serves two purposes- firstly to reduce the product temperature and
secondly to
enable partial separation of the naphtha component in the produced hydrocarbon
product, enriching the remaining liquid in the diesel component.
Figure 2 shows a process diagram, for flash separation of naphtha and diesel
hydrocarbons, in which:
1 is a fixed-bed Fischer Tropsch reactor.
2 is a mixture of gases, water, naphtha, diesel and light waxes at
ca.190-240 C and pressure greater than atmospheric.
3 is a pressure let-down valve.
4 is stream 2 at reduced temperature due to gas expansion
and at
14,7 psia.
5 is a flash drum vessel.
6 is a vapour phase consisting of steam 2 minus diesel and
light
waxes.
14
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
7 is a cooler.
8 is stream 6 with naphtha and water in the liquid phase.
9 is a vessel to retain naphtha and water.
is a waste tailgas stream consisting mainly of inert gases and
5 light hydrocarbons.
The FT products 2 flow through a pressure let-down valve 3 and into a flash
drum 5. The inert gases and lower-boiling hydrocarbons, water and naphtha go
overhead as vapour out of the flash dram and through cooler 7. The diesel and
light
10 waxes collect in vessel 5. The water and naphtha condense in cooler 7
and are
collected in vessel 9. The remaining gases exit overhead in stream 10 and are
typically combusted, sometimes with energy recovery, or are used to generate
electricity.
EXAMPLES
Catalyst Supports Employed
Table 1. Physical characteristics of catalyst supports
Alumina catalyst Alcoa Alcoa Alcoa F-220 Sasol
negussa
support (supplier's 1,13-5 CSS450 Trilobes Aerolyst
3038
__________ data) . _¨
Surface Area, tn2/g 300 min 350 360 248 270
Average Particle size, various 2120 2000 1670x4100
2500
microns
Pore Volume, cc/g 0763 0.57 0,5 0,82 0.9-1.0
"
Bulk Density, gice 0.645 0,72 0.769 0.42 0.40-0.46
A1203, %wt Diff. 99.6 93,1 Dff <500 PPIrt
S102,-lkowt 0.40 max 0,02 0.02 0.015 >99.8
Fe203, %wt, max 0.04 0.02 0.015 <30 ppm
Na20, %wts MEM 0,2 0,35 0.3 0.05
LOI (250-1100 C), %wt 23-30 3.5 6.5
Example 1
Catalyst synthesis was conducted by ordinary means as practiced by those
knowledgeable in the art. The catalyst support was alumina trilobe extrudate
obtained
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
from Sasol Germany GmbH (hereafter referred to as trilobe'). The extrudate
dimensions were 1.67 mm diameter and 4.1 mm length. The support was calcined
in
air at 500 C. for 24 hours. A solution mixture of cobalt nitrate and
perrhenie acid VMS
added to the support by the method of incipient wetness to achieve 5 wt%
cobalt
metal and 0.5 wt.% rhenium metal in the finished catalyst (Catalyst 1). The
catalyst
was oxidized in three steps:
Step 1: the catalyst was heated to 85 C and held for 6 hours.
Step 2: the temperature was raised to 100 C at 0.5 C per minute and held for
4-hours;
Step 3: the temperature was raised to 350 C at 0.3 C per minute and held for
12 hours,
The drying rate of the wet catalyst was somewhat dependent upon the size of
catalyst particles. Smaller particles will dry more quickly than larger
particles and the
size of the crystals formed inside the pores can vary with crystallization
rate. A
volume of 29 cc of oxidized catalyst was placed in a '/2 inch OD tube that had
an outer
annular space through which temperature-control water was flowed under
pressure in
order to remove the heat of reaction. In effect, the FT reactor was a shell-
and-tube
heat exchanger with catalyst placed in the tube side. The inlet gas and water
were both
at the targeted reaction temperature, Catalyst reduction was accomplished by
the
following procedure:
Reduction- gas flow rate (cc/min)/H2 in nitrogen (%)/temperature ( C.)/time
(hours):
1. 386/70/200/4, pre-heat stage
2, 386/80/to 325/4, slow heating stage
3, 386/80/325/30, fixed- temperature stage
During Fischer-Tropsch catalysis, total gas flow to the FT reactor was at a
OHSV
of 1000 VI. Gas composition was representative of an air-autothemial reformer
gas;
50% nitrogen, 33.3% H2 and 16.7 % CO, A seasoning of the catalyst was used to
16
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
reduce methane production. This was accomplished by holding the reactor
temperature at 170 C. for the first 24 hours. Presumably, this process causes
carbonylation of the cobalt surface and increased PT activity. CO conversion
and
liquid production were measured at a variety of temperatures between 190 C.
and
220 C.
Example 2
The catalyst used in this example (Catalyst 2) was the same as the catalyst
used in
Example 1, except that the cobalt metal loading was 10 wt%.
Example 3
The catalyst used in this example (Catalyst 3) was the same as the catalyst
used in Example 1, except that the cobalt metal loading was 15 wt /0.
Example 4
The catalyst used in this example (Catalyst 4) was the same as the catalyst
used in Example 1, except that the cobalt metal loading was 20 wt%.
Example 5
The catalyst used in this example (Catalyst 5) was the same the catalyst used
in Example 1, except that the cobalt metal loading was 26 wt%.
Example 6
The catalyst used in this example (Catalyst 6) was the same the catalyst used
in Example 1, except that the cobalt metal loading was 35 wt%.
Example 7
The catalyst used in this example (Catalyst 7) was the same as the catalyst
used in Example 1, except that the alumina support was CSS-350, obtained from
Alcoa, and the cobalt loading WU 20 weight percent. This support is spherical
with a
diameter of 1/16 inch.
17
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
pamples 8a1 Sb, Sc & 8d
The catalysts used in these examples (Catalysts 8a, 8b, 8c, and 8d) were the
same as used in Example 1, except as follows: Catalyst 8a was identical to
Catalyst 1,
except that the alumina support was LD-5, obtained from Alcoa, and the cobalt
loading was 20 weight percent. This support is spherical with an average
particle
distribution of 1963 microns. Example 8a used the particle size mixture as
received.
Some of the original particles were ground to smaller sieve sizes: Catalysts
8b, Sc and
84 were made with particles of diameter 214, 359 and 718 microns respectively.
The
cobalt loading in Examples 8b, So and 84 was identical to Catalyst 8a.
Example 9.
The catalyst used in this example (Catalyst 9) was the same as the catalyst
used in Example 1, except that the alumina support was F-220, obtained from
Alcoa,
and the cobalt loading was 20 weight percent. F-220 is a spherical support
with a
mesh size distribution of 7/14.
Example 10
The catalyst used in this example (Catalyst 10) was the same as Catalyst 4,
except that the promoter was ruthenium rather than rhenium.
Examnle 11
The catalyst used in this example (Catalyst 11) was the same as Catalyst 3,
except that Aerolyst 3038 silica catalyst support from Degussa was used
instead of
alumina.
Example 12
The catalyst used in this example (Catalyst 12) was identical with Catalyst 8d
having the same catalyst support, particle size and catalyst loading, except
that the
oxidizing process hold times were doubled during catalyst synthesis, That is,
the
temperature hold times were respectively to 12, 8 and 24 hours for the 3-
oxidizing
steps described for Catalyst 1. The intention of slower catalyst oxidation
rates of the
small Catalyst 12 particles was to achieve a larger cobalt crystallite size
(21.07 rim)
18
CA 02748216 2011-06-20
WO 2010/071989 PCT/CA2009/001862
within the pores of the small support particle in comparison with the
crystallite size
under faster crystallization conditions of Catalyst 8d (15.72). The method
used herein
to control drying rate and catalyst cobalt crystallite size is not meant to
exclude any
other method to achieve larger crystallite sizes. For example, the relative
humidity or
pressure of the drying chamber could be varied to control the catalyst drying
rate and
therefore cobalt crystallite size.
Catalyst Characterization
The above Catalysts were analyzed for average crystallite size (d(Co0),
Dispersion (D%) and Degree of Reduction (30R) using a Chembet 5000
(Quantachrome Instruments) TPR/TI'D analyzer. The catalyst was reduced at 325
C
in H2 flow and cobalt dispersion was calculated assuming that one hydrogen
molecule
covers two cobalt surface atoms. Oxygen chemisorption was measured with a
series
of (02/He) pulses passed through the catalyst at 380 C temperature after
reducing the
catalyst at 325 C. The up-take oxygen moles were determined and degree of
reduction was calculated assuming that all cobalt metal was re-oxidized to
Co304.
Cobalt crystallite size was calculated from:
d(Co0).-- (96/D%) DOR
D% ; Dispersion
14T Catalyst Evaluation
(I) Influence of cobalt loading
The effect of Co loading on catalyst performance was tested with Examples 1-
6 with the results shown in Table 2.
Table 2. Effect of catalyst loading on performance on Examples 1.6 (trilobes)
at 70 psia.
Weight % Cobalt
(Example number) 5 (1) 10(2) 15(3) 20 (4)
26(5) 36 (6)
Optimum Temperature, C 220 210 205 200 200 200
Hydrocarbon Liquid Rate, mith 0,09 0.54 0.74 1.03 0,77 0.86
Naphtha, wt% 8,4 8.8 13.9 17,9 18.4 15.8
Diesel, wt% 92,5 82.8 78,3 76.3 76.8 76.8
Light wax, wt% 1.1 8.4 7.8 6,9 6.8 7.4
19
CA 02748216 2011-06-20
WO 2010/071989 PCT/CA2009/001862
Diesel production, ml/h 0,08 0.48 0.58 0.78 0,59
0.66
CO Conversion, mol % 19.4 42.0 61,2 88.1 82.8
83.1
C5+ Selectivity, % 28,6 80,6 71,3 66.0 65,1
84.3
Octane number 81 79 77 75 74 75
Tests for each of Examples 1-6 were conducted at various temperatures, and
the temperature that gave the largest amount of hydrocarbon product is listed.
It is
clear that 5% cobalt was not enough to provide a useful amount of liquid
hydrocarbons: the best concentration was 20 wt% Co, which gave 1.03 ml/h. The
concentration of diesel range hydrocarbons in the hydrocarbon product was 75.3-
92.5
% at cobalt loadings of 10 wt% cobalt or higher. The highest diesel production
rate
(0.78 ml/h) was achieved with the trilobe support with 20% cobalt at 70 psia.
The performance data for Catalyst 1 at 202.5 is shown in Table 8. The level of
wax (C>23) on the C5+ liquid was only 6.8 % and the diesel fraction was 73,5%
(C9-
C23). It was found that for all Catalysts tested where the crystallite average
diameter
was greater than 16 nm, the C5+ wax was less than 10 weight %, enabling the
product
to be used directly as diesel blend.
Figure 3 shows the carbon number distribution for Catalyst 3 (trilobe) in
Example 3 at 190 C. A very narrow distribution was obtained having no heavy
wax.
Diesel was 90.8%, naphtha 6.1% and light waxes 3.1%. Cetane number was very
high at 88. In all graphs of carbon numbers, naphtha is indicated by large
squares,
diesel by diamonds and light waxes by small squares,
Influence of pressure
Catalyst 4 in Example 4 was run in the standard testing rig as described above
at a temperature of 202.5 C. at a variety of pressures. Results in Table 3
and Figure 4
indicate that productivity of the catalyst for production of liquid
hydrocarbons was
significant at low pressures down to 70 psia, with the optimum results
obtained at
pressures between 70 psia and 175 psia. Preferred pressures are 70-450 psia
and most
preferably from 70 to 175 psia. The diesel fraction over that pressure range
was fairly
constant at 70.8-73.5 weight percent. As shown in Table 8, Catalyst 4, with 20
%
CA 02748216 2011-06-20
WO 2010/071989 PCT/CA2009/001862
cobalt had an average crystallite size of 22.26 nm and a C5+ wax fraction of
6.8 wt %
enabling the product to be used as a diesel blend.
Table 3. Effect of pressure on catalyst performance (Catalyst 4, 202.5 C.).
Pressure, psis 40 70 100 125 140 175
200
Hydrocarbon Liquid Rate, mini 0.405
1.047 1.082 1.034 1.048 1.079 0,805
Naphtha, wt% 8.5 19.7 24,7 23.5 23,9
28.6 23,9
Diesel, wt% 77.8 73.5 71,9 73.1 73.4
70,8 74.1
Light wax, wt% 13,7 6.8 3.4 3,4 2,7 2.6
2,0
Diesel production, mVh 0.32 0.77 0.78 0,76 0,77
0.78 0.60
CO Conversion, mol % 59.4 90.2 84,1 83.8 74.8 73.4
65,8
C5+ Selectivity, % 78.6 58,1 54.4 52.5 61,3
57.7 52.0
Catalyst 7
As seen in Table 4, the maximum diesel production rate was achieved at 215
C. and 70 psia. Compared with Catalyst 4, Catalyst 7 gave a lower diesel
production
rate at its optimum temperature (215 C.), but a higher diesel fraction.
Figure 5 shows
the narrow carbon number range in the liquid product at 190 C., with 89.6%
in the
diesel range. Cetane number was 81. However, as shown in Table 8, the
crystallite
size was 18.26 urn, and the wax fraction was 7.2% enabling the product to be
used as
a diesel blend.
Table 4. Performance of Catalyst 7 at various temperatures (CSS-350).
Temperature, C 190 200 210 215 220
Hydrocarbon Liquid Rate, ml/h 0.55 0,58 0.64 0.70 0,68
Naphtha, wt% 5,4 15,2 13,4 16.4 14.3
Diesel, wt% 89,6 76.8 82,0 77.4 81.3
Light wax, wt% 5.0 8.0 4.8 7.2 4.4
Diesel production, mitt) 20.1 47,2 45.2 53.8 49.5
Average Molecular Weight 194.9 170.2 171.2 184.8
168.3
CO Conversion, mol % 47.8 53.4 81.6 93.8 100.0
Catalysts 8a. 8b. Sc. and Sd
The testing results are shown in Table 5, Catalysts 8b, Sc and 8d showed Co
metal dispersion higher than for Catalyst 8a. Catalysts that contain Co
average
crystallite sizes below 16 nanometers gave a high wax cut in the FT product of
17.6-
21
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
19,3% wt, whereas Catalyst 8a and Catalyst 12, which contained Co
crystallites
larger than 16 tun gave lower wax cuts of 6.6 and 7.8 wt.% respectively in the
C5+
liquid, enabling the product to be used as a diesel blend. Of note, Catalysts
8a and 12
had very different particle sizes, but gave similar low wax cuts. This shows
that the
controlling variable for low wax concentrations was crystallite size, and not
particle
size.
Table 5. Performance of Catalyst 82-84 and 12 at 70 psia.
Catalyst 8a 12 8b* 8c* 8d*
Average Particle size, microns 1963 718 274 359 718
Average crystallite size, nm 23.08 21.07 9.19 14.78
15.72
Dispersion, % 4.16 4.58 10.45 6.5 6,11
Fischer-Tropsch test Temperature, C 200 205 200 200 200
C5+ composition wt. %:
Naphtha (carbon number C6-C8) 9.3 16,7 10.1 10,4 11.4
Diesel (carbon number C9-023) 84.1 76,5 70.9 72 69.3
Wax (carbon number > C23) 6.6 7.8 19,0 17,6 19.3
Average molecular weight, AMU 182 160 195 190 193
CO conversion, % 62.7 51.2 88,8 72.7 69.3
*Not part of the present Application
to
catalvst
Catalyst 9 was tested at 70 psis, As shown in Table 6 and Figure 7, the 190 C
hydrocarbon product contained 99.1% "naphtha plus diesel". Diesel itself was
at
93.6%, There was very little light wax. Cetane number was 81. As shown in
Table 8,
the crystallite size was 22,22 nm and the wax fraction was 2.3 %, enabling the
product
to be directly as a diesel fuel.
Table 6. Performance of Catalyst 9 (F.220) at various temperatures (pressure
70
psis).
Temperature, C 190 200 210 215
Hydrocarbon Liquid Rate, mVh 0.485 0.757 0.8 0.733
Naphtha, wt% 5.5 9.2 20.1 21.5
Diesel, wt% 93,8 88.8 77.0 74.7
Light wax, wt% 0.9 2.3 2.9 3.8
Diesel production, ml/h 0.41 0.62 0.53 0.47
Average Molecular Weight 188,2 181,4 157,7 154.1
CO Conversion, mol % 50,0 72.2 94.7 92.2
Cetane number 81,0 76.0 67.0 65.0
22
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
Catalyst 10:
Data in Table 7 and Figure 8 show that the use of ruthenium catalyst
promoter
instead of rhenium also provides a narrow distribution of hydrocarbons with
74.42%
in the diesel range having an overall cetane number of 78. As shown in Table
8, the
crystallite size was 20.89 nm and the wax fraction was 3,73 %, enabling the
product
to be used as a diesel blend.
to
Table 7. Performance of Catalyst 10 (Ruthenium promoters LD-5 alumina
support).
Temperature, C/Pressure, psia 215
Conversion, % 94,84
C5+ liquid rate, ml/h 0,73
Diesel production rate, ml/h 0,54
C5+ weight Fractions, %:
Naphtha (C8-C8) 21,86
Diesel (C9-C23) 74,42
Wax (>023) 3,73
Cetane number 78
Average molecular weight 184
Catalut 11
For Catalyst 11, the hydrocarbon liquid production rate was 0.55 ml/h at 210
C. The carbon distribution curve shown in Figure 9 demonstrates a narrow
distribution with a high diesel cut. As shown in Table 8. the crystallite size
was 33.1
rim and the wax fraction was 5.2 %, enabling the product to be used as a
diesel blend,
perhaps after flashing off the naphtha fraction.
Table 8. Summary of the effect of cobalt crystallite size on C5+ wax
concentration.
Catalyst Number 4 7 9 10 11 12
Name Trilobes CSS- F-220 LID-
Aerolyst LD-5
350 5/Ru 3038
Average crystallite size, nm 22.26 18.28 22,22 20,89 23,10
21.07
Dispersion, % 4.31 5.28 4.32 4.60 2.90
4.58
F-T test temperature, C 202.5 215 200 215 200
205
23
CA 02748216 2011-06-20
WO 2010/071989
PCT/CA2009/001862
C5+ composition wt. %;
Naphtha (carbon number C6-C8) 10.7 15,4 9.2 21.9 20.7 157
Diesel (carbon number C9-C23) 73.5 77.4 88,6 74.4 74.1
76.5
Wax (carbon number > C23) 6.8 7.2 2.3 3.7 5.2 7.8
Average molecular weight, AMU 165 165 181 164 147 180
CO conversion, % 90,2 93.8 722 92.8 67 51.2
Catalysts 1 to 12 (except catalysts 8 b, c and d) in this disclosure show that
a
narrow distribution of hydrocarbons, mainly in the diesel range, having low
wax
content (<10 wt.%) is obtained when the FT catalyst has cobalt crystallites
larger than
16 nm, as shown in Figure 10 (the large squares are not part of this
embodiment).
With small catalyst particles (e.g. Catalyst 12) it is necessary to control
the
crystallization rate in order to obtain the desired crystallite size.
Figure 11 compares this result with expectations from the Anderson-Shultz-
Flory (A-S-F) carbon number distribution based on chain growth. The A-S-F
distribution provides only 50 wt. % diesel fraction, whereas the present
embodiments
provide >65 wt. %.
The liquid hydrocarbon product of the present catalysts is more valuable than.
the broad A-S-F type of product because it can be used directly as a diesel-
blending
stock without hydrocracking to increase cetame number and decrease sulphur
content
of petroleum diesels. Because the present process can be a simple once-through
process, it can entail low capital cost.
Although the disclosure describes and illustrates preferred embodiments of
the invention, it is to be understood that the invention is not limited to
these particular
embodiments, Many variations and modifications will now occur to those skilled
in
the art. For a complete definition of the invention and its intended scope,
reference is
to be made to the summary of the invention and the appended claims read
together
with and considered with the disclosure and drawings herein.
24
